The present invention relates to a process for the production of acetic acid by the carbonylation of methanol and/or a reactive derivative thereof in the presence of a Group VIII noble metal catalyst and a hydrocarbyl halide co-catalyst.
Processes for producing acetic acid by the Group VIII noble metal catalysed, hydrocarbyl halide co-catalysed carbonylation of alcohols and/or their reactive derivatives are well-known in the art. Representative of such art employing rhodium as the Group VIII noble metal catalyst may be mentioned, for example, U.S. Pat. No. 3,772,380; GB-A-1468940; GB-A-1538783 and EP-A-0087070. Representative of such art using iridium as the Group VIII noble metal catalyst may be mentioned, for example, GB-A-1,234,121; U.S. Pat. No. 3,772,380; DE-A-1767150; EP-A-0616997; EP-A-0618184; EP-A-0618183; and EP-A-0657386.
In continuous liquid phase processes for the production of acetic acid by the carbonylation of methanol and/or a reactive derivative thereof in the presence of a Group VIII noble metal the acetic acid product is recovered from the liquid reaction composition and dried; the remaining components of the reaction composition being recycled to the reactor to maintain their concentration therein.
Howard et al in Catalysis Today, 18(1993), 325-354 describe the rhodium and iridium catalysed carbonylation of methanol to acetic acid. The continuous rhodium-catalysed, homogeneous methanol carbonylation process is said to consist of three basic sections; reaction, purification and off-gas treatment. The reaction section comprises a stirred tank reactor, operated at elevated temperature and pressure, and a flash vessel. Liquid reaction composition is withdrawn from the reactor and is passed through a flashing valve to the flash tank where the majority of the lighter components of the liquid reaction composition (methyl iodide, methyl acetate and water) together with product acetic acid are vapourised. The vapour fraction is then passed to the purification section whilst the liquid fraction (comprising the rhodium catalyst in acetic acid) is recycled to the reactor (as shown in FIG. 2 of Howard et al). The purification section is said to comprise a first distillation column (the light ends column), a second distillation column (the drying column) and a third distillation column (the heavy ends column) (as shown in FIG. 3 of Howard et al). In the lights ends column methyl iodide and methyl acetate are removed overhead along with some water and acetic acid. The vapour is condensed and allowed to separate into two phases in a decanter, both phases being returned to the reactor. Wet acetic acid is removed from the light ends column typically as a side draw and is fed to the drying column where water is removed overhead and an essentially dry acetic acid stream is removed from the base of the distillation zone. From FIG. 3 of Howard et al it can be seen that the overhead water stream from the drying column is recycled to the reaction section. Heavy liquid by-products are removed from the base of the heavy ends column with product acetic acid being taken as a side stream.
In practice the upper (aqueous layer) from the decanter, in whole or in part, is returned to the light ends column as reflux and the lower (organic layer) from the decanter is recycled to the reactor. For operational reasons it is highly desirable that two separable phases are maintained in the decanter. Decanter stability is of paramount importance in the successful operation of the continuous carbonylation process. If the decanter becomes single phase, the resulting composition change tends to increase the water content in the reactor, which in turn has a significant impact on reaction activity for iridium catalysed carbonylation.
EP-A-0768295 describes one method of maintaining two separable phases in the reactor in circumstances such that the concentration of water contained in the carbonylation liquid reaction composition decreases or the concentration of methyl acetate contained in the liquid reaction composition increases. Thus EP-A-0768295 discloses a process for producing acetic acid by reacting continuously at least one selected from methanol, methyl acetate and dimethyl ether with carbon monoxide in the presence of a Group VIII metal-containing catalyst, methyl iodide and water, comprising (a) a step in which a crude reaction liquid is withdrawn from a carbonylation step and introduced into a flash zone, and a catalyst circulating liquid containing a catalyst component which is not evaporated in the flash zone is circulated into a carbonylation reactor, (b) a step in which a vapour fraction evaporated in the flash zone is fed into a first distillation column in the form of a vapour or a liquid, (c) a step in which a low boiling circulating stream comprising water, methyl acetate, methyl iodide and acetic acid is withdrawn from the top of the first distillation column, and (d) a step in which crude acetic acid is withdrawn from the bottom or the side cut near the bottom of the first distillation column, characterised in that a liquid separation state in the decanter at the top of the first distillation column is maintained by adding water to the first distillation column, lowering the cooling temperature at the overhead part of the first distillation column, or reducing the concentration of methyl acetate contained in the liquid fed into the decanter at the top of the first distillation column.
EP-A-0768295 teaches that when two phases do not form in the decanter liquid and the unseparated liquid is recycled to the reactor, by-product carbonyl compounds, such as acetaldehyde, crotonaldehyde and 2-ethylcrotonaldehyde, and organic iodine compounds such as hexyl iodide, build up to an unacceptable level in the product acetic acid.
European patent publication number EP-0573189-A1 describes a process for the production of acetic acid by carbonylation of methanol in the presence of a rhodium carbonylation catalyst. The methyl acetate concentration in the liquid reaction composition is said to be at least 2% by weight, preferably in the range 2% to 15% by weight more preferably in the range 3% to 10% by weight. Whilst in Examples 4 and 5 the combined overhead streams forming the light ends recycles shown were calculated to have 0.96% and 1.33% by weight acetic acid, the methyl acetate concentrations in the reactors were only 3.1% and 7.3% by weight.
We have found that at high methyl acetate concentrations, typically 8% w/w or greater in the liquid reaction composition in the carbonylation reactor, particularly at low levels of water and methyl iodide, which conditions are typically associated with the use of iridium as the carbonylation catalyst, it becomes increasingly difficult to achieve two separable phases in the decanter, which in turn may give rise to product quality problems of the type referred to in EP-A-0768295, and plant capacity problems, largely as a result of hydraulic limitations to both control valves and pumps.
We have found that a solution to the problem of maintaining two liquid phases in a continuously operated decanter is to control the concentration of acetic acid in the overhead fraction fed from the light ends column to the decanter. EP-A-0768295 makes no mention of acetic acid concentration in the overhead fraction and its impact on the maintenance of two phases. In off-line experiments we have found that a typical decanter feed will form a single phase with about 14% w/w or more of acetic acid present. However, in a continuously operated decanter, even lower levels of acetic acid must be achieved (8 wt % or lower) in order to maintain stable operation. This is due to the increasing water content of the organic phase, which depletes the light ends column overheads of water by recycling it directly back to the reactor. This causes the water concentration to fall and the phase separation to become more difficult. A feed-back mechanism then becomes dominant and the decanter becomes single phase.
Accordingly the present invention provides a continuous process for the production of acetic acid by the carbonylation of methanol and/or a reactive derivative thereof which process comprises the steps of:
(I) feeding methanol and/or a reactive derivative thereof to a carbonylation reactor in which the methanol and/or reactive derivative thereof is reacted with carbon monoxide in a liquid reaction composition, the liquid reaction composition comprising a Group VIII noble metal carbonylation catalyst, methyl iodide co-catalyst at a concentration of at least 2% w/w, optionally at least one promoter, at least a finite concentration of water, methyl acetate at a concentration of at least 8% w/w and acetic acid product;
(II) withdrawing liquid reaction composition from the carbonylation reactor and introducing the withdrawn liquid reaction composition into at least one flash separation zone, with or without the addition of heat, to produce a vapour fraction comprising water, acetic acid product, methyl acetate and methyl iodide, and a liquid fraction comprising Group VIII noble metal carbonylation catalyst and optionally at least one promoter;
(III) recycling the liquid fraction from step (II) to the carbonylation reactor;
(IV) introducing the vapour fraction from step (II) into a light ends distillation column;
(V) removing a process stream comprising acetic acid product from the light ends distillation column;
(VI) removing from the head of the light ends distillation column a vapour fraction comprising methyl acetate, methyl iodide, water and acetic acid;
(VII) condensing the overhead vapour fraction from (VI);
(VIII) passing the condensed overhead vapour fraction from (VII) to a decanter wherein the fraction is separated into an upper (aqueous) layer and a lower (organic) layer;
(IX) recycling in whole or in part the upper (aqueous) layer separated in (VIII) as reflux to the light ends distillation column and the lower (organic) layer separated in (VIII) in whole or in part to the reactor characterised in that separability of an upper (aqueous) layer and a lower (organic) layer in the decanter in step (VIII) is achieved by maintaining the concentration of acetic acid in the condensed overhead vapour fraction passed to the decanter at or below 8 wt %.
The concentration of acetic acid in the condensed vapour fraction passed to the decanter is preferably maintained below 8 wt %, preferably below 6 wt %, more preferably less than 5 wt %. Maintenance of the concentration of acetic acid in the condensed vapour fraction within the aforesaid ranges is largely achievable by suitable operation of the light ends distillation column. Thus, the reflux ratio within the column and/or the number of theoretical stages in the column are selected such that the acetic acid concentration in the condensed vapour fraction is 8 wt %, or below. Typically, the light ends column contains a relatively small number of stages (around 10 in total). It has been found that the aqueous phase must all be refluxed to the column to maintain two liquid phases in practice in a commercial unit operating with about 10 theoretical stages above the feed. It is preferred that the light ends column has greater than 10, more preferably 15, or greater, theoretical stages above the feed. Increasing the number of theoretical stages allows lower reflux ratios to be employed, which gives a benefit in terms of water removal efficiency and thus reduced purification costs. Another modification by which the acetic acid concentration in the decanter may be maintained within the aforesaid limits is to relocate any recycle streams having a substantial acetic acid content, which otherwise may formerly have been fed to the condenser and thus directly into the decanter, to the light ends distillation column, suitably at a point close to the feed point of the vapour fraction from step (II) so as to allow the acetic acid in the recycle stream to be separated out from this stream by the stages above the feed. Such a recycle stream may be, for example, a vapour return stream from the off-gas treatment section of the process.
As regards the decanter itself, a conventional design for methanol carbonylation plants includes the provision of a boot, which takes the form of a short vertical cylindrical section depending from the horizontal cylindrical section. This is a standard design feature for systems where there is either a low volume flow of heavy phase, or where the heavy phase density is very high and it is desirable to minimise the inventory of heavy phase material. It has been found that under the relatively high methyl acetate concentration conditions prevailing in the process of the present invention it is possible to eliminate the boot normally present in the construction of the decanter. Elimination of the boot from the decanter provides the advantage of capital cost savings due to simpler fabrication of the decanter vessel. It also avoids the possibility of poorer separation caused by turbulence within the boot induced by high volume flows.
It is further preferred that the decanter contains plate pack separators, which are commercially available (from, for example Natco, Tulsa, Okla.), to enhance the rate of phase separation. Plate pack separators generally comprise stacks of inclined, corrugated plates which induce coalescence and reduce the residence time required in the decanter. Installation of plate pack separators has the advantage that it facilitates the use of smaller decanters. In turn this leads to the advantage that if the decanter becomes single phase, the disadvantageous impact of increased water content in the reactor referred to hereinabove is minimised.
In step (I) of the process of the present invention methanol and/or a reactive derivative thereof is fed to a carbonylation reactor. Suitable reactive derivatives of methanol include methyl acetate and dimethyl ether.
The methanol and/or reactive derivative thereof is reacted in the carbonylation reactor with carbon monoxide in a liquid reaction composition. The carbon monoxide may be essentially pure or may contain inert impurities such as carbon dioxide, methane, nitrogen, noble gases, water, and C1 to C4 paraffinic hydrocarbons. The presence of hydrogen in the carbon monoxide feed and generated in situ by the water gas shift reaction is preferably kept low as its presence may result in the formation of hydrogenation products. Thus, the amount of hydrogen in the carbon monoxide reactant is preferably less than 1 mol %, more preferably less than 0.5 mol % and yet more preferably less than 0.3 mol % and/or the partial pressure of hydrogen in the carbonylation reactor is preferably less than 1 bar partial pressure, more preferably less than 0.5 bar and yet more preferably less than 0.3 bar. The partial pressure of carbon monoxide in the reactor is suitably in the range greater than 0 to 40 bar, typically from 4 to 30 bar.
The liquid reaction composition in the reactor comprises a Group VIII noble metal carbonylation catalyst, methyl iodide co-catalyst optionally at least one promoter, at least a finite concentration of water, methyl acetate at a concentration of at least 8% w/w and acetic acid product.
Of the Group VIII noble metals rhodium and iridium are preferred. The noble metal catalyst may comprise any metal-containing compound which is soluble in the liquid reaction composition. The metal catalyst may be added to the liquid reaction composition in any suitable form which dissolves in the liquid reaction composition or is convertible therein to a soluble form. Suitable compounds are described in the aforesaid patent publications relating to iridiumxe2x80x94and rhodium catalysed carbonylations. Typically carbonyl complexes, halide salts and acetate salts of the metals may be employed. Rhodium may be present in an amount of from 50 to 5000 ppm, preferably from 100 to 1500 ppm. Iridium may be present in an amount in the range from 100 to 6000 ppm, preferably from 400 to 3000 ppm.
As co-catalyst there is used methyl iodide. Methyl iodide may suitably be present in the liquid reaction composition in an amount in the range from 2 to 20%, preferably from 4 to 16% by weight.
The choice of promoter when present in the liquid reaction composition depends to some extent on the nature of the Group VIII noble metal catalyst. When iridium is employed as the carbonylation catalyst the optional promoter is suitably a metal selected from the group consisting of ruthenium, osmium, cadmium, rhenium, mercury, gallium, indium, tungsten, and mixtures thereof, preferably ruthenium or osmium. Suitably the molar ratio of promoter: iridium is in the range [0.5 to 15]:1. When rhodium is employed as the carbonylation catalyst the optional promoter is suitably selected from the group consisting of iodide salts of alkali and alkaline earth metals, for example lithium iodide, quaternary ammonium iodides, and quaternary phosphonium iodides. Suitably the optional promoter may be present up to its limit of solubility.
Irrespective of the Group VIII noble metal used as carbonylation catalyst the liquid reaction composition in the carbonylation reactor contains at least a finite concentration of water. However, the amounts of water may vary depending on the Group VIII noble metal employed as catalyst. Generally, for rhodium water may be present in an amount in the range from 0.1 to 30%, preferably from 1 to 15% by weight. For iridium water may be present in an amount from 0.1 to 10%, preferably from 1 to 6.5% by weight.
Methyl acetate, irrespective of whether or not it is fed to the carbonylation reactor, is inevitably present in the liquid reaction composition by reason of the reaction of methanol and/or a reactive derivative thereof with acetic acid present as the carbonylation product and/or carbonylation solvent. Insofar as the present invention is concerned methyl acetate is present in the liquid reaction composition in an amount of 8 wt % or greater, typically 8 to 50 wt %, preferably 8 to 35 wt %. Generally, these methyl acetate concentration ranges are those associated with iridium as the Group VIII noble metal catalyst, the methyl acetate concentration using rhodium as catalyst generally, but not necessarily, being at the most 5 wt %, typically below about 3 wt %.
The remainder of the liquid reaction composition comprises acetic acid.
The carbonylation reaction temperature is suitably in the range from 100 to 300xc2x0 C., preferably in the range from 150 to 220xc2x0 C. The total pressure in the carbonylation reactor is suitably in the range from 10 to 200 barg, preferably 15 to 100 barg, more preferably 15 to 50 barg.
In step (II) of the process of the present invention liquid reaction composition is withdrawn from the carboniylation reactor and introduced into at least one flash separation zone, with or without the addition of heat, to produce a vapour fraction comprising water, acetic acid product, methyl acetate and methyl iodide, and a liquid fraction comprising Group VIII noble metal carbonylation catalyst and optionally at least one promoter. If a single stage flash is used the pressure may be in the range 0 to 3 barg, with a temperature suitably in the range 100 to 150xc2x0 C. Using a two-stage flash, the pressure in the first flash may be in the range 1 to 10 barg and the pressure in the second flash may suitably be in the range 0 to 5 barg.
In step (III) of the process the liquid fraction recovered from the flash separation zone in step (II) is recycled to the carbonylation reactor.
In step (IV) of the process the vapour fraction recovered from the flash separation zone in step (II) is introduced into a light ends distillation column. Suitably, the light ends distillation column has up to 40 theoretical stages. The column may be operated at any suitable pressure, for example a heads pressure of about 1.2 barg and a base pressure of about 1.5 barg. The operating temperature of the light ends distillation column will depend upon a number of factors, including the composition of the feed, heads and base streams and the operating pressure. Typical base temperatures may be in the range 125 to 140xc2x0 C. and typical head temperatures may be in the range 105 to 115xc2x0 C.
In step (V) of the process a stream comprising acetic acid product is removed from the light ends distillation column. The process stream may be removed at any suitable point, for example above or below the feed point, or as a liquid or vapour from the base of the column. The process stream comprising acetic acid product removed from the light ends distillation column may then be dried, for example, in a drying distillation column, the separated water suitably being either recycled to the carbonylation reactor or removed from the process. The dried acetic acid may suitably then be passed to a heavy ends distillation column in which propionic acid by-product is separated from dry acetic acid.
In step (VI) of the process a vapour fraction comprising methyl acetate, methyl iodide, water and acetic acid is removed from the head of the light ends distillation column.
In step (VII) of the process the overhead vapour fraction from (VI) is condensed.
In step (VIII) of the process the condensed overhead fraction from (VII) is passed to a decanter wherein the fraction is separated into an upper (aqueous) layer and a lower (organic layer).
Finally, in step (IX) of the process the upper (aqueous) layer separated in (VIII) is recycled in whole or in part as reflux to the light ends distillation column and the lower (organic) layer separated in (VIII) is recycled in whole or in part, preferably in whole, to the reactor. The tipper (aqueous) layer is suitably returned in part to the light ends distillation column as reflux, suitably at a rate of about 0.1 to about 0.7 times the rate of removal of the vapour fraction from the head of the light ends distillation column.